Abstract:

The present invention provides a reactor system having: (1) a first
reactor receiving an oxygenate component and a hydrocarbon component and
capable of converting the oxygenate component into a light olefin and the
hydrocarbon component into alkyl aromatic compounds; (2) a separator
system for providing a first product stream containing a C3 olefin,
a second stream containing a C7 aromatic, and a third stream
containing C8 aromatic compounds; (3) a first line connecting the
separator to the inlet of the first reactor for conveying the second
stream to the first reactor; (4) a second line in fluid communication
with the separator system for conveying the C3 olefin to a propylene
recovery unit, and (4) a third line in fluid communication with the
separator system for conveying the C8 aromatic compounds to a xylene
recovery unit.

Claims:

1. A reactor system comprising:a first reactor having an inlet, an outlet
and a reaction zone, the first reactor is capable of receiving a feed
stream having an oxygenate component and a hydrocarbon component and of
converting a portion of the oxygenate component to a light olefin and
alkylating a portion of the hydrocarbon component to form alkyl aromatic
compounds, the light olefin and the alkyl aromatic compounds are included
in a first effluent stream;a separator system in fluid communication with
the outlet of the first reactor for receiving the first effluent stream
of the first reactor and for providing a first product stream containing
a C3 olefin, a second stream containing a C7 aromatic, and a
third stream containing C8 aromatic compounds;a first line
connecting the separator to the inlet of the first reactor for conveying
the second stream to the first reactor;a second line in fluid
communication with the separator system for conveying the C3 olefin
to a propylene recovery unit; anda third line in fluid communication with
the separator system for conveying the C8 aromatic compounds to a
xylene recovery unit.

2. The reactor system of claim 1 further comprising a second reactor for
receiving a C.sub.6.sup.+ aromatic hydrocarbon containing feed stream
from the separator system and subjecting the feed stream to conditions
for transalkylating the C.sub.6.sup.+ aromatic hydrocarbon to form an
effluent stream rich in C8 hydrocarbons.

3. The reactor system of claim 1 further comprising a third reactor for
receiving a C.sub.4.sup.+ olefin stream from an outlet of the separator
and forming an effluent stream rich in C.sub.3.

4. The reactor system of claim 3 wherein the effluent stream rich in
C3 from the third reactor is in fluid communication with an inlet of
the separator system.

5. The reactor system of claim 3 further comprising a catalyst
regeneration unit in fluid communication with both the first reactor and
the third reactor.

6. The reactor system of claim 1 wherein the oxygenate component is
selected from the group consisting of methanol and dimethyl ether.

7. The reactor system of claim 1 further comprising a first catalyst
chamber in the first reactor for guiding a flow of a catalyst material
through a portion of the first reactor.

8. The reactor system of claim 1 further comprising a steam stripper in
fluid communication with the first line for concentrating the aromatic
components of the second feed stream.

Description:

CROSS-REFERENCE TO RELATED APPLICATIONS

[0001]This application is a Division of prior copending application Ser.
No. 11/955,610 filed Dec. 13, 2007, which is incorporated herein by
reference in its entirety and which claims the benefit of U.S.
Provisional Application No. 60/878,204 filed Dec. 29, 2006.

FIELD OF THE INVENTION

[0002]The present invention relates generally to a reactor system and a
process for producing aromatics and particularly xylenes in conjunction
with a methanol-to-propylene reactor system.

BACKGROUND OF THE INVENTION

[0003]A major portion of the worldwide petrochemical industry is concerned
with the production of light olefin materials and their subsequent use in
the production of numerous important chemical products via
polymerization, oligomerization, alkylation and the like well-known
chemical reactions. Light olefins include ethylene, propylene and
mixtures thereof. These light olefins are essential building blocks for
the modern petrochemical and chemical industries. The major source for
these materials in present day refining is the steam cracking of
petroleum feeds. For various reasons including geographical, economic,
political and diminished supply considerations, the art has long sought a
source other than petroleum for the massive quantities of raw materials
that are needed to supply the demand for these light olefin materials.
Thus, R & D personnel seek to use alternative feedstocks effectively and
selectively to produce light olefins, thereby lessening dependence of the
petrochemical industry on petroleum feedstocks. Much attention has been
focused on the possibility of using hydrocarbon oxygenates and more
specifically methanol or dimethylether (DME) as a prime source of the
necessary alternative feedstock. Oxygenates are particularly attractive
because they can be produced from such widely available materials as
coal, natural gas, recycled plastics, various carbon waste streams from
industry and various products and by-products from the agricultural
industry. The art of making methanol and other oxygenates from these
types of raw materials is well established and typically involves the use
of one or more of the following procedures: (1) manufacture of synthesis
gas by any of the known techniques typically using a nickel or cobalt
catalyst in a steam reforming step followed by the well-known methanol
synthesis step using relatively high pressure with a copper-based
catalyst; (2) selective fermentation of various organic agricultural
products and by-products in order to produce oxygenates; or (3) various
combinations of these techniques.

[0004]Given the established and well-known technologies for producing
oxygenates from alternative non-petroleum raw materials, the art has
focused on different procedures for catalytically converting oxygenates
such as methanol into the desired light olefin products in order to make
an oxygenate to olefin (OTO) process. These light olefin products that
are produced from non-petroleum based raw materials must of course be
available in quantities and purities such that they are interchangeable
in downstream processing with the materials that are presently produced
using petroleum sources. Although many oxygenates have been discussed in
the prior art, the principal focus of the two major routes to produce
these desired light olefins has been on methanol conversion technology
primarily because of the availability of commercially proven methanol
synthesis technology. Two principal techniques are known in the art for
conversion of methanol to light olefins (MTO). U.S. Pat. No. 4,387,263
discloses one MTO processes that utilizes a catalytic conversion zone
containing a zeolitic type of catalyst system. The '263 patent reports on
a series of experiments with methanol conversion techniques using a ZSM-5
type of catalyst system.

[0005]U.S. Pat. No. 4,587,373 discloses using a zeolitic catalyst system
like ZSM-5 for purposes of making light olefins. The '373 patent
discloses diverting a portion of a methanol feed stream to a DME
absorption zone to allow for downsizing of a scrubbing zone.

[0006]U.S. Pat. No. 5,095,163; U.S. Pat. No. 5,126,308 and U.S. Pat. No.
5,191,141 disclose an MTO conversion technology utilizing a non-zeolitic
molecular sieve catalytic material. More particularly these patents
disclose using a metal aluminophosphate (ELAPO) and more specifically a
silicoaluminophosphate molecular sieve (SAPO) and even more specifically
SAPO-34. This SAPO-34 material was found to have a very high selectivity
for light olefins with a methanol feedstock and consequently very low
selectivity for the undesired corresponding light paraffins and the
heavier materials.

[0007]The classical OTO technology produces a mixture of light olefins
primarily ethylene and propylene along with various higher boiling
olefins. Although the classical OTO process technology possesses the
capability of shifting the major olefin product recovered therefrom from
ethylene to propylene by various adjustments of conditions maintained in
the reaction zone, the art has long sought an oxygenate to propylene
(OTP) technology that would provide better yields of propylene relative
to the classical OTO technology. The driving force for this shift in
emphasis towards propylene is the growth rate of the propylene market
versus the growth rate of the ethylene market. The existing sources of
propylene production in the marketplace are primarily based on
conventional steam cracking of naphtha, LPG streams, propane streams and
the like. Another principal source of propylene is produced in a fluid
catalytic cracking (FCC) hydrocarbon conversion process in the modern day
refinery.

[0008]US 2003/0139635A1 discloses a fixed bed methanol to propylene (MTP)
process for selectively producing propylene from a feedstock of methanol
and/or DME. This patent application discloses a flowscheme having an
oxygenate to propylene (OTP) synthesis portion having three reactors in a
parallel flow arrangement with respect to the oxygenate feed and utilize
a steam diluent and fixed beds of oxygenate conversion catalysts. The
reactors are connected in a serial flow arrangement with respect to the
effluents of the first reactor and the second reactor.

[0009]EP-B-1025068 discloses using two reaction zones to convert an
oxygenate feed and a by-product fraction containing C4.sup.+
hydrocarbons to ethylene and propylene. This patent discloses that the
two reaction zones allow for independent selection of catalyst and
conversion conditions for each zone. This patent discloses using a
non-zeolitic molecular sieve catalyst such as SAPO-34 for an oxygenate to
light olefin reaction zone and either a non-zeolitic molecular sieve
catalyst or a zeolitic catalyst such as ZSM-5 material for the auxiliary
reaction zone which operates to convert the C4.sup.+ by-product
fraction to the desired light olefin (i.e., C2 and C3 olefins).
The patent discloses using a circulating fluid bed or a riser reaction
for the first reaction zone and a fluid bed or a fixed bed or a fixed
tube reactor for the second reaction zone.

SUMMARY OF THE INVENTION

[0010]The present invention provides a reactor system having: (1) a first
reactor having an inlet, an outlet and a reaction zone, the first reactor
is capable of receiving a feed stream having an oxygenate component and a
hydrocarbon component and of converting a portion of the oxygenate
component to a light olefin and alkylating a portion of the hydrocarbon
component to form alkyl aromatic compounds, the light olefin and the
alkyl aromatic compounds are included in a first effluent stream; (2) a
separator system in fluid communication with the outlet of the first
reactor for receiving the first effluent stream of the first reactor and
for providing a first product stream containing a C3 olefin, a
second stream containing a C7 aromatic, and a third stream
containing C8 aromatic compounds; (3) a first line connecting the
separator to the inlet of the first reactor for conveying the second
stream to the first reactor; (4) a second line in fluid communication
with the separator system for conveying the C3 olefin to a propylene
recovery unit, and (4) a third line in fluid communication with the
separator system for conveying the C8 aromatic compounds to a xylene
recovery unit.

[0011]The present invention further provides a process for converting an
oxygenate feed to propylene (C3 olefin) and a hydrocarbon feed to
alkyl substituted benzene compounds. The process includes: (1) reacting
in a reaction zone the oxygenate feed in the presence of a molecular
sieve catalyst to convert a portion of oxygenate compounds in the
oxygenate feed to a C3 olefin; (2) reacting in the reaction zone the
hydrocarbon feed in the presence of the molecular sieve catalyst to form
C6.sup.+ compounds; (3) forming an effluent stream containing the
C3 olefin and the C6.sup.+ compounds; (4) separating a portion
of the effluent stream in a separation zone into a first stream rich in
the C3 olefin, a second stream rich in C7; (5) passing the
first stream to a first storage unit; and (6) recycling the second stream
to an inlet of the reaction zone.

[0012]The present invention further provides a process for preparing
xylenes. The process includes: (1) providing a first stream containing
aromatic compounds under pressure to a first reactor, the first reactor
being operated under conditions for converting an oxygenate compound to a
C3 olefin, the first reactor containing a molecular sieve type
catalyst; (2) alkylating a portion of the first stream in the first
reactor to form an effluent stream rich in C6.sup.+ aromatic
compounds; (3) separating a portion of the effluent stream in a
separation zone into a first stream rich in a C7 aromatic compound
and into a second stream rich in C8 aromatic compounds; (4)
recycling the first stream to an inlet of the first reactor; (5) and
passing the second stream to a recovery unit.

BRIEF DESCRIPTION OF THE DRAWING

[0013]The FIGURE is a schematic view of a reactor system for recycling
C7 aromatic compounds to an oxygenate to propylene reactor system
and for producing and isolating C8 aromatic compounds (xylenes).

TERMS AND CONDITIONS DEFINITIONS

[0014]The following terms and conditions are used in the present
specification with the following meanings (1) A "portion" of a stream
means either an aliquot part that has the same composition as the whole
stream or a part that is obtained by eliminating a readily separable
component therefrom (e.g., if the stream contains hydrocarbons in
admixture with steam, then after condensation of a major portion of the
steam, it comprises an aqueous portion and a hydrocarbon portion); (2)
the presence of necessary compressors and/or pumps is understood when
flow is shown from a zone of relatively low pressure to a zone of higher
pressure; (3) the presence of necessary heating and/or cooling means is
implied when flow is shown between zones operating at different
temperatures; (4) the term "light olefins" means ethylene, propylene and
mixtures thereof; (5) the term "heavy olefin" means an olefin having a
molecular weight greater than propylene; (6) the expression "OTP" process
means a process for converting an oxygenate to propylene and in a
preferred embodiment when the oxygenate is methanol the OTP process is
referred to as an "MTP" process herein; (7) the term "oxygenate" means an
oxygen-substituted aliphatic hydrocarbon containing 1 to 10 carbon atoms
include aliphatic alcohols, ethers, and carbonyl compounds (e.g.,
aldehydes, ketones, carboxylic acids, and the like) and mixtures of these
materials; (8) the term "highly unsaturated hydrocarbon" means a
hydrocarbon which contains two or more double bonds or a triple bond in
its structure; (9) the term "fluidized bed" means particles of a catalyst
are entrained in a pressurized stream of gas or liquid; (10) the term
Cx.sup.+ refers to a hydrocarbon compound having x number of carbon
atoms or greater and the number x can equal anywhere from 2 to 30
carbons; and (11) the term "purge" and as referenced in the FIGURE is for
removing water, paraffins and unwanted or unreactive compounds as is well
known to those skilled in the art.

DETAILED DESCRIPTION OF THE INVENTION

[0015]The following description of a preferred embodiment of the process
of the present invention is made with reference to the attached FIGURE.
In the interest of simplifying the description of the invention in order
to facilitate understanding the FIGURE do not contain representations of
heaters, heat exchangers, coolers, valves, control means and other
conventional items that are well known to those of ordinary skill in the
chemical engineering art except where their presence is essential to the
understanding of the present invention.

[0016]The FIGURE shows a reactor system 10 for co-producing xylenes with
propylene from an oxygenate-containing hydrocarbon feed stream 11 in an
OTP reactor 14 having an effluent stream 15. The reactor system 10 has a
separator unit having a first stage 18 for separating C6.sup.+
aromatic hydrocarbons from light and heavy olefins. The C6.sup.+
aromatic hydrocarbons are part of an effluent stream 20 and are directed
through an optional steam stripper 21, which concentrates the
C6.sup.+ aromatic hydrocarbons and removes a portion of the heavy
olefins overhead, and the heavy olefins are removed and transferred
through line 25 to an olefin interconversion reactor 62. From the steam
stripper, line 23 carries the C6.sup.+ aromatic hydrocarbons to a
second stage 22 where the C6.sup.+ aromatic hydrocarbons are
separated in a distillation column or a series of distillation columns
into a first effluent stream 24 rich in C7, a second effluent stream
26 rich in C8, a third effluent stream 28 rich in C6 and a
fourth effluent stream 29 rich in C9.sup.+ aromatic compounds. A
portion of the C9.sup.+ aromatic compounds in line 29 may be purged
in line 36. The first effluent stream 24 is recycled back to join stream
17 and form part of the feed stream 11. The second effluent stream 26 is
directed to a xylene recovery unit 30. The third effluent stream 28
merges with the fourth effluent stream 29 and the combined flow is
directed to a second reactor 40 for performing transalkylating reactions.

[0017]The light and heavy olefins form part of an effluent stream 50 that
is directed from the separator 18 after compression to an olefin
separation unit 52 where C3 olefin (propylene) is separated from
lighter and heavier olefins and passed through effluent line 54 to a
propylene recovery unit 56. The C2 olefin forms part of an effluent
line 58. C2 olefin product can be recovered in line 34. The
remaining C2 olefin is recycled back in stream 59 to join stream 17
and form part of the feed stream 11. The C4.sup.+ olefins
(C4-C8) are passed through line 60 from which a portion of the
C4.sup.+ olefins may be recovered through line 35. The remaining
C4.sup.+ olefins are passed through line 61 to join heavy olefins in
line 25 via line 63. The heavy olefins and C4.sup.+ olefins in line
63 are passed to a third reactor which is the olefin interconversion
reactor 62 in which a portion of the C4.sup.+ olefins are converted
to C3 olefins which form part of an effluent stream 64 and are
directed back to and inlet of the separator 18. Water is recycled back to
the first reactor 14 through lines 57 and 56 and a portion of the water
can also be purged through line 33.

[0018]The first reactor 14 converts oxygenate to propylene and is known in
the art as an OTP reactor. In a preferred form of the OTP reactor 14, the
oxygenate-containing hydrocarbon feed is catalytically and selectively
converted to propylene and by-product hydrocarbons containing aliphatic
moieties such as--but not limited to--methane, ethane, ethylene, propane,
butylene, butane and limited amounts of other higher carbon number
aliphatics by contacting the feedstock with an OTP catalyst at effective
OTP conditions. The oxygenate feed stream 11 will contain, in a preferred
form of the invention, one, some or all of methanol, dimethyl ether
(DME), ethanol, diethyl ether, methylether, formaldehyde, dimethyl
ketone, acetic acid, and mixtures thereof. In a most preferred form of
the invention, the oxygenate feed stream 11 will contain methanol or
dimethylether or mixtures thereof. A fresh oxygenate feed stream 12 can
be first conveyed to an optional reactor 13 for converting at least a
portion of the methanol to dimethyl ether and an effluent stream 17 of
the oxygenate feed stream is then conveyed with other streams into the
first reactor 14 via the oxygenate feed stream 11.

[0019]A diluent is not absolutely required in the first reactor 14 but is
a useful option to maintain the selectivity of the OTP catalyst to
produce light olefins, particularly propylene. The use of a diluent such
as steam can provide certain equipment cost and thermal efficiency
advantages as well as lowering the partial pressure of the oxygenate
reactants, thereby increasing selectivity to olefins. The phase change
between steam and liquid water can also be employed to advantage in
transferring heat between the feedstock and the reactor effluent, and the
separation of the diluent from the product requires only a simple
condensation step to separate water from the light olefin products.

[0020]A diluent is thus preferably used in the OTP reactor 14 to control
partial pressure of the oxygenate reactant to provide a heat sink for the
net exothermic reactions occurring therein and to shift the overall
reaction selectivity towards propylene. Suitable diluents for use in the
reaction zones include helium, argon, nitrogen, carbon monoxide, carbon
dioxide, hydrogen, water, C1-C5 paraffins, aromatic
hydrocarbons and mixtures of these materials. Preferred diluents are
steam, methane, an aromatic compounds, and mixtures thereof. Preferred
diluents are relatively inert at the conditions maintained in the
reaction zones. An especially preferred diluent is steam since it is
relatively easily recovered from the effluent stream utilizing
condensation techniques. The amount of diluent used will be selected from
the range of 0.1:1 to 12:1 and more typically from about 0.1:1 to 5:1
moles of diluent per mole of oxygenate in order to lower the partial
pressure of the oxygenates to a level which favors production of
propylene. In a preferred form of the present invention and as described
above, the reactor 14 will be supplied with a portion of an ethylene-rich
by-product stream 58 and a C7 rich stream 24. The C2 olefin
recycle stream and the C7 rich stream 24 will thus furnish a
hydrocarbon diluent to a reaction zone of the reactor 14 and therefore
the amount of diluent that must be added to achieve the target diluent to
oxygenate mole ratio will diminish once the reaction zone is started up
and C2 and the C7 rich stream 24 recycle stream is initiated.

[0021]The conversion conditions used in the reactor 14 is carefully chosen
to favor the production of propylene from the oxygenate components of the
oxygenate-containing hydrocarbon feed stream 11. In a preferred form of
the invention, oxygenate conversion temperatures will be from about
350° to about 600° C. The lower portion of this oxygenate
conversion temperature range with certain catalysts is known to favor the
production of propylene with the upper portion favoring the production of
ethylene at the expense of propylene. Preferred inlet temperatures into
the reaction zones are therefore in the range of 350° to
500° C., more preferably in the range of about 375° to
500° C. and most preferably in the range of 375° to
475° C.

[0022]These reaction conditions have been found to also be effective in
alkylating aromatic compounds to form C8 compounds (xylenes). Thus,
the oxygenate-containing hydrocarbon feed stream further includes
aromatic components C5.sup.+ that are alkylated to form C6 and
C6.sup.+ hydrocarbon compounds, and, therefore alkylated aromatic
compounds can be produced at the same time that oxygenates are being
converted to propylene.

[0023]These reactions are carried out in the presence of a catalyst and
more preferably a molecular sieve catalyst. The catalyst is contacted
with the feed stream 11 in the reactor using a fluidized bed, moving bed
or batch type catalyst distribution systems. In one preferred form of the
invention the reactor 14 will contain a moving bed catalyst system.

[0024]Suitable catalysts include zeolitic molecular sieves in the calcined
form be represented by the general formula:

Me2/nO:Al2O3:xSiO2:yH2O

where Me is a cation, x is the framework SiO2 to Al2O3
ratio and has a value from about 2 to infinity, n is the cation valence
and y has a value of about 2 to 100 or more and more typically about 2 to
25.

[0025]Zeolites which may be used include chabazite--also referred to as
Zeolite D, clinoptilolite, erionite, ferrierite, mordenite, Zeolite A,
Zeolite P, ZSM-5, ZSM-11, and MCM-22. Zeolites having a high silica
content (i.e., those having framework silica to alumina ratios greater
than 100 and typically greater than 150 with good results achieved at a
silica to alumina mole ratio of about 150:1 to 800:1) are especially
preferred. One such high-silica-content zeolite having the structure of
ZSM-5 is silicalite, as the term used herein includes both the
silicapolymorph disclosed in U.S. Pat. No. 4,061,724 and also the
F-silicate disclosed in U.S. Pat. No. 4,073,865. Best results are
obtained with ZSM-5 or ZSM-11 or a mixture thereof. Such catalyst are
sometimes referred to as having a "pentasil-type" structure.

[0026]Suitable non-zeolitic molecular sieves are embraced by an empirical
chemical composition, on an anhydrous basis, expressed by the empirical
formula:

(ELXAlyP.sub.Z)O2

where EL is an element selected from the group consisting of silicon,
magnesium, zinc, iron, cobalt, nickel, manganese, chromium and mixtures
thereof, x is the mole fraction of EL and is at least 0.005, y is the
mole fraction of aluminum and is at least 0.01, z is the mole fraction of
phosphorous and is at least 0.01 and x+y+z=1. When EL is a mixture of
metals, x represents the total amount of the element mixture present.
Preferred elements (EL) are silicon, magnesium and cobalt with silicon
being especially preferred.

[0027]It is contemplated using blends of zeolitic-type catalyst, blends of
non-zeolitic molecular sieve catalysts and blends of both zeolitic-type
and non-zeolitic-type molecular sieve catalyst.

[0028]The second reactor 40 is operated under conditions to transalkylate
C6 (benzene) and C6.sup.+ aromatic compounds and under
conditions to optimize the production of C8 aromatics. The
C6.sup.+ aromatic compounds include benzene, alkyl substituted
benzenes including methyl benzene (toluene), dimethyl benzenes (xylenes),
trimethyl benzenes, tetramethyl benzenes, and C2-C6 alkyl
substituted benzenes. The C8 aromatics are directed to a first
effluent stream 26 and are passed to the xylene recovery unit 30. The
C8 aromatics include o-xylene, m-xylene and p-xylene and most
preferably p-xylene. The second reactor 40 forms a second effluent stream
46 rich in C7 aromatics which are recycled back to the separator 22.

[0029]To effect a transalkylation reaction, the present invention
incorporates a transalkylation catalyst in at least one zone, but no
limitation is intended in regard to a specific catalyst other than size
and shape. Conditions employed in the transalkylation zone normally
include a temperature of from about 200° to about 540° C.
The transalkylation zone is operated at moderately elevated pressures
broadly ranging from about 100 kPa to about 6 MPa absolute. The
transalkylation reaction can be effected lover a wide range of space
velocities, with higher space velocities effecting a higher ratio of
para-xylene at the expense of conversion. Weighted hourly space velocity
(WHSV) generally is in the range of from about 0.1 to about 20 hr-1.

[0030]The transalkylation effluent is separated into a light recycle
stream, a mixed C8 aromatics product and a heavy-aromatics stream.
The mixed C8 aromatics product can be sent for recovery of
para-xylene and other valuable isomers. The light recycle stream may be
diverted to other uses such as to benzene and toluene recovery, but
alternatively is recycled partially to the transalkylation zone. The
heavy recycle stream contains substantially all of the C9 and
heavier aromatics and may be partially or totally recycled to the
transalkylation reaction zone.

[0031]One skilled in the art is familiar with several types of
transalkylation catalysts that may be suitably sized and shaped for use
in the present invention. For example, in U.S. Pat. No. 3,849,340, which
is herein incorporated by reference, a catalytic composite is described
comprising a mordenite component having a SiO2/Al2O3 mole
ratio of at least 40:1 prepared by acid extracting Al2O3 from
mordenite prepared with an initial SiO2/Al2O3 mole ratio
of about 12:1 to about 30:1 and a metal component selected from copper,
silver and zirconium. Friedel-Crafts metal halides such as aluminum
chloride have been employed with good results and are suitable for use in
the present process. Hydrogen halides, boron halides, Group I-A metal
halides, iron group metal halides, etc., have been found suitable.
Refractory inorganic oxides, combined with the above-mentioned and other
known catalytic materials, have been found useful in transalkylation
operations. For instance, silica-alumina is described in U.S. Pat. No.
5,763,720. Crystalline aluminosilicates have also been employed in the
art as transalkylation catalysts. ZSM-12 is more particularly described
in U.S. Pat. No. 3,832,449. Zeolite beta is more particularly described
in Re. 28,341. A favored form of zeolite beta is described in U.S. Pat.
No. 5,723,710. The preparation of MFI topology zeolite is also well known
in the art. In one method, the zeolite is prepared by crystallizing a
mixture containing an alumina source, a silica source, an alkali metal
source, water and an alkyl ammonium compound or its precursor. Further
descriptions are in U.S. Pat. No. 4,159,282; U.S. Pat. No. 4,163,018 and
U.S. Pat. No. 4,278,565.

[0032]A refractory binder or matrix is optionally utilized to facilitate
fabrication of the catalyst, provide strength and reduce fabrication
costs. The binder should be uniform in composition and relatively
refractory to the conditions used in the process. Suitable binders
include inorganic oxides such as one or more of alumina, magnesia,
zirconia, chromia, titania, boria, thoria, phosphate, zinc oxide and
silica. Alumina is a preferred binder.

[0033]The catalyst also contains an optional metal component. One
preferred metal component is a Group VIII (IUPAC8-10) metal,
preferably a platinum-group metal. Alternatively a preferred metal
component is rhenium. Of the preferred platinum group, i.e., platinum,
palladium, rhodium, ruthenium, osmium and iridium, and platinum is
especially preferred. This component may exist within the final catalytic
composite as a compound such as an oxide, sulfide, halide, or oxyhalide,
in chemical combination with one or more of the other ingredients of the
composite, or, preferably, as an elemental metal. This component may be
present in the final catalyst composite in any amount which is
catalytically effective, generally comprising about 0.01 to about 2
mass-% of the final catalyst calculated on an elemental basis. The
platinum-group metal component may be incorporated into the catalyst in
any suitable manner such as coprecipitation or cogellation with the
carrier material, ion exchange or impregnation. Impregnation using
water-soluble compounds of the metal is preferred. Typical platinum-group
compounds which may be employed are chloroplatinic acid, ammonium
chloroplatinate, bromoplatinic acid, platinum dichloride, platinum
tetrachloride hydrate, tetraamine platinum chloride, tetraamine platinum
nitrate, platinum dichlorocarbonyl dichloride, dinitrodiaminoplatinum,
palladium chloride, palladium chloride dihydrate, palladium nitrate, etc.
Chloroplatinic acid is preferred as a source of the especially preferred
platinum component. Moreover, when the metal component is rhenium,
typical rhenium compounds which may be employed include ammonium
perrhenate, sodium perrhenate, potassium perrhenate, potassium rhenium
oxychloride, potassium hexachlororhenate (IV), rhenium chloride, rhenium
heptoxide, and the like compounds. The utilization of an aqueous solution
of perrhenic acid is highly preferred in the impregnation of the rhenium
component Rhenium may also be used in conjunction with a platinum-group
metal.

[0034]The third reactor 62, in a preferred form of the invention, is an
interconversion reactor capable of converting non-propylene hydrocarbons
having fewer than three carbon atoms or four carbon atoms or greater, for
example C4-C8 hydrocarbons, into propylene The third reactor 62
preferably utilizes the same catalyst as used in the first reactor 14 and
contacts feed stream 60 to the catalyst using a moving bed, fluidized
bed, or batch type reactor systems. Also, in a preferred form of the
invention, the first reactor and the third reactor share a catalyst
regenerator system 70 connected to the first reactor through line 72 and
the third reactor through line 74. Regenerated catalyst lines 73 and 75
return regenerated catalyst to reactors 14 and 62, respectively. An
effluent from the third reactor 62 is transferred through the line 64 to
an inlet of the separator 18 via line 19.

[0035]A diluent may be used in the interconversion reactor 62 to control
the partial pressure of the heavy olefin reactant used therein and to
provide an additional heat source for the endothermic interconversion
reaction. Suitable diluents can be chosen from those previously set forth
in connection with the operation of the OTP reaction zones. Of these
preferred diluents, steam involves the risk of hydrothermal deactivation
of the catalyst used in the interconversion reactor if steam is used in
high concentration but is typically used because of its ability to
control and/or prevent coke formation in heaters, heat exchangers and
reactor internals, its ready availability, its ease of separability from
the products of the interconversion reaction and because it can be used
at a much lower concentration than in the OTP reaction zones. The amount
of diluent preferably used in the interconversion reaction zone
corresponds 0.001:1 to 1:1 moles of diluent per mole of C4.sup.+
olefin charged to this zone and more preferably to a mole ratio of 0.01:1
to 0.5:1. Unlike the situation with respect to the OTP reaction zones it
is to be noted that since H2O is not a by-product of the
C4.sup.+ interconversion reactions performed in the interconversion
reactor, there is typically no net make of diluent across this zone so
that the effective amount of diluent used in the interconversion reactor
is the amount charged thereto. However, it is within the scope of the
present invention to charge some oxygenate to the interconversion reactor
in an amount sufficient to off-set the endothermic interconversion
reactions arising therein.

[0036]One preferred form of the present invention utilizes moving bed
technology in the OTP reactor 14 and the olefin interconversion reactor
62 to enhance the selectivity of the overall process for propylene
production. The use of moving bed technology in a classical MTO process
is known and is shown in U.S. Pat. No. 5,157,181.

[0037]Moving bed reaction zones for use in the instant invention can be
configured in a number of ways, for example, the catalyst particles can
be introduced to an upper section of the OTP reaction zones and fed by
gravity through the entire volume of the reaction zones, wherein the
dual-function catalyst is contacted, in a preferred form of the
invention, a radially flowing feed stream; thus, the fluid stream or
streams flow transversely to the direction of flow of the catalyst. It is
contemplated, that the feed streams or by-product stream could be
directed to flow in a countercurrent direction to the catalyst movement
or in a concurrent direction without departing from the scope of the
present invention.

[0038]More typically the catalyst particles are introduced into an annular
catalyst chamber, or annular catalyst chambers, defined by concentric
catalyst retaining screens that run through the reactors wherein the
catalyst particles travel down through the annular catalyst chamber and
are withdrawn from a lower section of these reaction zones.

[0039]During the traversal through the reactors, a carbonaceous material,
i.e., coke, is deposited on the catalyst as it flows through the
reactors. The carbonaceous deposit material has the effect of reducing
the number of active sites on the catalyst which thereby affects the
extent of the overall conversion and the selectivity to propylene. A
portion of the coked catalyst is thus withdrawn from the reactors and
regenerated in station 70 to remove at least a portion of the coke
therefrom. In a preferred form of the invention where the same catalyst
is used in the first reactor 14 and the third reactor 62 coked particles
from both reactors can be mixed together and charged to the common
regeneration station 70. It is within the scope of the present invention
to charge at least a portion of the partially coked catalyst particles
withdrawn from the third reactor to the first OTP reactor. This can be
advantageous when the selectivity of the catalyst to propylene in the
first reactor is improved due to the partial coverage of active sites
with fresh coke deposits.

[0040]The carbonaceous material is removed from the catalyst by oxidative
regeneration wherein a moving bed of the catalyst particles withdrawn
from the reactors is contacted with an oxygen-containing gas stream at
sufficient temperature and oxygen concentration to allow the desired
amount of the carbonaceous materials to be removed by combustion from the
catalyst.

[0041]Both the oxygenate to propylene conversion and the C4.sup.+
olefin interconversion steps are effectively carried out over a wide
range of pressures including inlet total pressures between about 0.1 atm
(10.1 kPa) up to about 100 atm (10.1 MPa) but it is well known that the
formation of lighter olefins like propylene are favored at low pressure
conditions. It is thus preferred for both of these steps to use an inlet
pressure in the range of about 1 to 4 atm (101.3 to 405 kPa) and best
results are achieved at about 1.4 to about 3.4 atm (138 to 345 kPa).

[0042]The contact time of the reactants with the catalyst is ordinarily
measured in relative terms of a Weight Hourly Space Velocity (WHSV) which
is calculated for the OTP conversion step on the basis of mass hourly
flow rate of the sum of the mass of oxygenate reactants passed to the OTP
reaction zone plus the mass of any reactive hydrocarbon material present
in the feed stream or any of the recycle streams passed to the first
reaction zone divided by the mass of the dual-function catalyst present
in the reaction zone. The WHSV for the C4.sup.+ olefin
interconversion step is likewise calculated on the basis of mass hourly
flow rate of the sum of the mass of C4.sup.+ olefin by-product
stream passed thereto plus the mass of any reactive hydrocarbons present
in any recycle stream or diluent stream passed thereto divided by the
mass of the catalyst present in the third reactor 62. Those skilled in
the art will recognize that the contact time of the reactants with the
catalyst is proportional to the inverse of the WHSV such that as the WHSV
increases contact time decreases and conversely a decrease in WHSV
produces an increase in contact time. WHSV for use in both the OTP
reactors and the interconversion reactor associated with the present
invention can range from about 0.1 to 100 hr-1, with a preferred
range being about 0.5 to 20 hr-1, with best results ordinarily
attained in the range of 0.5 to 10 hr-1.

[0043]The present invention further includes an optional selective
hydrogenation treatment step to selectively hydrogenate highly
unsaturated hydrocarbons such as dienes and/or acetylenic hydrocarbons
that are formed in the OTP conversion step in minor amounts (i.e., less
than 2 wt-% of the amount of oxygenate feed converted and typically about
0.01 to 1 wt-% of the amount converted). While these highly unsaturated
hydrocarbons do not represent a substantial source of propylene yield
loss, it has been found that they are a very significant contributor to
the rate of coke deposition on the preferred catalyst. The selective
hydrogenation conditions utilized in this treatment step are selected
from conditions known to those of skill in the art to be effective to
convert highly unsaturated hydrocarbons to the corresponding olefins
while minimizing or eliminating any over-hydrogenation to the
corresponding fully saturated hydrocarbon.

Example 1

[0044]A pilot plant test was conducted using a methanol and ethylene feed
stream having 30.4 g/hr methanol, 4.8 g/hr of ethylene and 15.0 g/hr of
water. The plant was operated at 440° C. at an inlet, under a
pressure of 76 kPa (11 psig) at the inlet a GHSV of 6,100 hr-1. The
feed stream was contacted with a ZSM-5 catalyst. A product was collected
from the plant and was found to have by weight on a dry basis 25.3%
propylene, 32.2% ethylene, 0.38% benzene, 0.38% toluene, 1.37% xylenes
and 0.68% of C9.sup.+.

Example 2

[0045]A second pilot plant test was conducted using a feed stream having
by weight 30.5 g/hr methanol, 15.0 g/hr; water, 4.2 g/hr ethylene and
50.9 g/hr heavy hydrocarbons. The heavy hydrocarbons comprised by weight
approximately 25% C5, 12.5% hexene, 25% toluene, 12.5% octenes and
25% trimethyl benzenes. Hence the combined feed contained 19.4% toluene
and 19.4% C9 aromatics by weight on a dry basis. The pilot plant
reactor was operated at 440° C. at an inlet, under a pressure of
122 kPa (17.7 psig) at the inlet a GHSV of 6,100 hr-1. ZSM-5
catalyst was used as in Example 1. An effluent from the reactor was
collected and analyzed to have by weight on a dry basis 12.2% propylene,
5.12% ethylene, 0.4% benzene, 13.6% toluene, 7.6% xylene and 22.3%
C9.sup.+. Example 2 shows an increased selectivity over Example 1 in
terms of xylene and C9 aromatics while showing a disappearance of
toluene.